Processes for recovering paraxylene

ABSTRACT

Disclosed is a process for recovering paraxylene in which a first simulated moving bed adsorption unit is used to produce two extract streams—one rich in paraxylene and a paraxylene-rich extract stream that is lean in ethylbenzene and an ethylbenzene-rich extract stream that is lean in paraxylene- and a paraxylene-depleted raffinate stream. A significant amount of the ethylbenzene is removed in the ethylbenzene-rich extract stream (at least enough to limit buildup in the isomerization loop), so the paraxylene-depleted raffinate stream may be isomerized in the liquid phase. Avoiding vapor phase isomerization saves energy and capital, as liquid phase isomerization requires less energy and capital than the vapor phase isomerization process due to the requirement of vaporizing the paraxylene-depleted stream and the use of hydrogen, which requires an energy and capital intensive hydrogen recycle loop.

PRIORITY CLAIM

This application claims priority to U.S. Provisional Application Ser.No. 62/356,634 filed Jun. 30, 2016 and European Application No.16185649.7 filed Aug. 25, 2016, the disclosures of which are fullyincorporated herein by their reference.

TECHNICAL FIELD

This application relates to an improved process for producingparaxylene, particularly the recovery of paraxylene from streamscontaining ethylbenzene.

BACKGROUND

Ethylbenzene (EB), paraxylene (PX), orthoxylene (OX), and metaxylene(MX) are often present together in C₈ aromatic product streams fromchemical plants and oil refineries. High purity EB is an important rawmaterial for the production of styrene; however, for a variety ofreasons all high purity EB feedstocks used in styrene production areproduced by alkylation of benzene with ethylene, rather than by recoveryfrom a C₈ aromatics stream. Of the three xylene isomers, PX has thelargest commercial market and is used primarily for manufacturingterephthalic acid and terephthalate esters for use in the production ofvarious polymers such as poly(ethylene terephthalate), poly(propyleneterephthalate), and poly(butene terephthalate). While OX and MX areuseful as solvents and raw materials for making products such asphthalic anhydride and isophthalic acid, market demand for OX and MX andtheir downstream derivatives is much smaller than that for PX.

Given the higher demand for PX as compared with its other isomers, thereis significant commercial interest in maximizing PX production from anygiven source of C₈ aromatic materials. However, there are a number ofmajor technical challenges to be overcome in achieving this goal ofmaximizing PX yield. For example, the four C₈ aromatic compounds,particularly the three xylene isomers, are usually present inconcentrations dictated by the thermodynamics of production of the C₈aromatic stream in a particular plant or refinery. As a result, the PXproduction is limited, at most, to the amount originally present in theC₈ aromatic stream unless additional processing steps are used toincrease the amount of PX and/or to improve the PX recovery efficiency.A variety of methods are known to increase the concentration of PX in aC₈ aromatics stream. These methods normally involve cycling the streambetween a separation step, in which at least part of the PX is recoveredto produce a PX-depleted stream, and a xylene isomerization step, inwhich the PX content of the PX-depleted stream is returned back towardsequilibrium concentration.

In a typical aromatics plant, such as that shown in FIG. 1, liquid feed,typically a C₈₊ aromatic feedstream which has previously been processedby known methods to remove C⁷⁻ species (particularly benzene andtoluene), is fed by conduit 1 to xylenes re-run 3, an apparatus per sewell known in the art. The xylenes re-run (or more simply afractionation column) 3 vaporizes the feed and separates the C₈aromatics into an overhead mixture 5 of xylenes and EB, and a bottomproduct 61 comprising C₉₊ aromatics. The overhead mixture typically hasa composition of about 40-50% MX, 15-25% PX, 15-25% OX, and 10-20% EB.Unless otherwise noted herein, percentages are % weight. The overhead isthen condensed in condenser 7, an apparatus also per se well-known inthe art, and becomes the feed for the PX recovery unit 15, via conduit 9and 13, a portion of the condensed overhead may be returned to re-run 3as reflux via conduits 9 and 11.

The PX recovery unit 15 may employ crystallization technology,adsorption technology, or membrane separation technology, each per sewell known in the art. These technologies separate PX from its isomersand are capable of producing high purity PX up to 99.9%, which is takenfrom unit 15 via conduit 17. Shown in FIG. 1 is the case where unit 15is an adsorptive separation unit, such as a Parex™ or Eluxyl™ unit, inwhich case typically the extract 17, which comprises a desorbent, suchas paradiethylbenzene (PDEB), needs to be separated, such as bydistillation, from the desired extract PX in distillation column 19,which generates an overhead 23 that is condensed in condenser 25 toyield a liquid stream 27, which is a high purity PX stream. This stream27 may be taken off via conduit 31 and optionally a portion may bereturned to column 19 as reflux via conduit 29. The desorbent isreturned to the PX recovery system 15 via conduit 21. Raffinate from therecovery system 15, comprising MX, OX, EB, and some PX, is removed viaconduit 65 and sent to unit 37, discussed below. Note: a portion ofraffinate in 65 may be recovered and marketed as low-value solventxylene.

The raffinate 65, which comprises mainly MX, OX, EB, and desorbent issent to fractionation column 37, generating overhead 33 and bottoms 63.Overhead 33 contains MX, OX and EB, which is condensed in condenser 32and sent via conduit 35 and then 41 to isomerization unit 43, discussedin more detail below. A portion may be returned to fractionator 37 viaconduit 35 and then 39 as reflux. The desorbent in the bottoms productis returned to 15. Note that as used herein the term “raffinate” is usedto mean the portion recovered from the PX recovery unit 15, whether thetechnology used is adsorptive separation, crystallization, or membrane,and then is sent to the isomerization unit 43, conventionally a vaporphase isomerization unit, which uses technology also per se well-known.

A stream consisting essentially of MX, OX, and EB is sent toisomerization unit 43, an apparatus per se known in the art, toisomerize the MX and OX and optionally EB to PX. Isomerization unit 43may be a vapor phase or liquid phase isomerization unit. Conventionallythere are one or more heat exchangers or furnaces associated with thesystem shown in FIG. 1 between the PX recovery unit 15 and theisomerization unit that are not shown for convenience of view. Likewise,hydrogen separators and hydrogen compressors are also not shown forconvenience of view. These and other features, such as valves and thelike, would be apparent to one of ordinary skill in the art.

The product of the isomerization unit 43 is sent via conduit 51 to theC⁷⁻ distillation tower 53, which separates the product of isomerizationinto a bottom stream 59 comprising equilibrium xylenes and the overhead47, comprising C⁷⁻ aromatics, e.g., benzene and toluene. The overheadproduct is condensed in condenser 45 and then the distribution of liquidproduct via conduit 49 may be apportioned as desired between conduit 57and conduit 55, the former of which may be disposed of in numerous wayswhich would be well-known per se in the art, and the latter conduitreturning C⁷⁻ aromatics as reflux to tower 53. The bottoms product 59 ofdistillation tower 53 is then sent to xylenes re-run 3, either mergingwith feed 1 as shown in the figure, or it may be introduced by aseparate inlet (not shown).

Vapor phase isomerization of the raffinate's EB is generally needed tolimit EB buildup in the recycle stream 59. However, vapor phaseisomerization has many disadvantages, including high energy consumption,costly and complex process equipment, and high xylenes loss due toconversion of the xylenes in the raffinate into undesirable productssuch as light gases and heavy aromatics, e.g., by one or moreside-reactions such as one or more of cracking, transalkylation, ordisproportionation. Attempts to overcome these disadvantages includereducing the quantity of raffinate going to the vapor phaseisomerization, e.g., removing EB from the raffinate by usingchromatographic EB separation in the PX recovery unit, as disclosed inU.S. Pat. No. 7,915,471 and U.S Patent Publication No. 2015/0266794. Theseparated EB is isomerized in a vapor phase isomerization stage, withthe remainder of the raffinate being isomerized in a liquid phaseisomerization stage. The liquid phase isomerization stage is operatedunder conditions which lessen undesired cracking, transalkylation, anddisproportionation side-reactions. Isomerates from the vapor phase andliquid phase isomerization stages are then combined and recycled to thexylene re-run column.

Even when EB is separated for vapor phase isomerization, with theremainder of the PX-depleted raffinate subjected to liquid phaseisomerization, the vapor phase isomerization stage contributes to xyleneloop inefficiencies. Some of these inefficiencies result from one ormore of (i) the need to vaporize the separated EB and then re-condensethe vapor phase isomerate for combining with the isomerate derived fromthe liquid phase isomerization stage, (ii) the need to separateunreacted molecular hydrogen vapor for re-use as an isomerization treatgas, and (iii) the need for removing non-aromatics formed duringisomerization. Consequently, it is desired to further lessen or eveneliminate the need for vapor phase isomerization.

BRIEF SUMMARY

Embodiments disclosed herein are directed to processes for recovering PXin which a first simulated moving bed adsorption unit is used to producetwo extract streams—a PX-rich extract stream that is lean in EB, and anEB-rich extract stream that is lean in PX—and a PX-depleted raffinatestream. A significant amount of the EB is removed in the EB-rich extractstream (at least enough to limit buildup in the isomerization loop), sothe PX-depleted raffinate stream may be isomerized in the liquid phase.Avoiding vapor phase isomerization saves energy and capital, as liquidphase isomerization requires less energy and capital than the vaporphase isomerization process due to the requirement of vaporizing thePX-depleted stream and the use of hydrogen, which requires an energy-and capital-intensive hydrogen recycle loop. Liquid phase isomerizationalso reduces the destruction of xylene molecules into less valuablenon-aromatics. PX and EB may be recovered from the EB-rich extractstream, or PX may be recovered and the remaining C₈ isomers may beisomerized in the vapor phase to produce more PX.

In some embodiments, the process comprises introducing a firsthydrocarbon feed stream, which comprises a mixture of PX, MX, OX, andEB, and a desorbent stream, which comprises desorbent, into a firstsimulated moving bed adsorption unit. A first PX-rich extract stream,which comprises desorbent and PX, is withdrawn from the first simulatedmoving bed adsorption unit at a location downstream of the desorbentintroduction location and upstream of the feed introduction location. Inaddition an EB-rich extract stream, which comprises desorbent, EB, andPX, is withdrawn from the first simulated moving bed adsorption unit ata location downstream of the PX-rich extract stream withdrawal locationand upstream of the feed introduction location. Further, a firstPX-depleted raffinate stream, which comprises desorbent, MX, OX, and EB,is also withdrawn from the first simulated moving bed adsorption unit.

At least a portion of the first PX-depleted raffinate stream isisomerized in the liquid phase to produce a first isomerized streamhaving a higher PX concentration than the first PX-depleted raffinatestream, and the first isomerized stream is recycled to the firstsimulated moving bed adsorption unit. The EB-rich extract stream may besent to a PX recovery unit to produce a second PX-rich extract streamand either recover a pure EB stream or isomerize the remaining C₈isomers isomerized in the vapor phase to produce more PX. In oneembodiment, the PX recovery unit is a second simulated moving bedadsorption unit. In another embodiment, the second PX recovery unit is acrystallizer.

These and other objects, features, and advantages of the disclosedembodiments will become apparent in the following detailed description,drawings, specific embodiments, experiments, and accompanying claims.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 illustrates a conventional PX production and extraction process.

FIG. 2 illustrates a PX production and extraction process in accordancewith at least some embodiments disclosed herein.

FIG. 3 illustrates another PX production and extraction process inaccordance with at least some embodiments disclosed herein.

FIG. 4A illustrates a conventional simulated moving bed adsorption unit.

FIG. 4B illustrates a simulated moving bed adsorption unit as used in atleast some of the embodiments disclosed herein.

FIGS. 5 and 6 depict the concentration profiles of each component insidea first simulated moving bed unit of at least some of the embodimentsdisclosed herein.

FIG. 7 illustrates a PX production and extraction process in accordancewith at least some embodiments disclosed herein.

DETAILED DESCRIPTION

As used herein the term “C_(n)” hydrocarbon, wherein n is a positiveinteger, means a hydrocarbon having n number of carbon atom(s) permolecule. For example, a C₈ aromatic hydrocarbon means an aromatichydrocarbon or mixture of aromatic hydrocarbons having 8 carbon atom(s)per molecule. The term “C_(n)+” hydrocarbon, wherein n is a positiveinteger, means a hydrocarbon having at least n number of carbon atom(s)per molecule, whereas the term “C_(n)−” hydrocarbon wherein n is apositive integer, means a hydrocarbon having no more than n number ofcarbon atom(s) per molecule.

Embodiments disclosed herein include improved processes for recoveringPX. With reference to FIGS. 2 and 3, a first hydrocarbon feed 102comprising xylenes and EB is provided to a first simulated moving bedadsorption unit 120, where a PX-rich extract stream 122, an EB-richextract stream 124, and a PX-depleted raffinate stream 126 arerecovered. The PX-depleted raffinate stream 126 is passed to a xyleneisomerization unit 140 where the PX-depleted raffinate stream 126 isisomerized under at least partial liquid phase conditions to produce anisomerized stream 142 having a higher PX concentration than thePX-depleted raffinate stream 126. The isomerized stream 142 is thenrecycled to the first PX recovery unit 120 to recover additional PX andthe process is repeated.

Optionally, the EB-rich extract stream 124 is further processed in a PXrecovery unit to recover additional PX. In one embodiment (e.g., seeFIG. 1), the PX recovery unit is a second simulated moving bedadsorption unit 220. When the second simulated moving bed adsorptionunit 220 is used, the EB-rich extract stream 124 may or may not passthrough a fractionation tower to remove the desorbent before the secondsimulated moving bed adsorption unit 220. The EB-rich extract stream 124may also be combined with a second hydrocarbon feed 202 prior to thesecond simulated moving bed adsorption unit 220. In another embodiment(e.g., see FIG. 2), the PX recovery unit is a crystallizer 300 and theEB-rich extract stream 124 may or may not pass through a fractionationtower to remove the desorbent before the crystallizer 300. In yetanother embodiment, the EB-rich extract stream 124 is sent to therefinery for further processing or for the motor gasoline pool.

The first hydrocarbon feed stream 102 employed in the present processmay be any hydrocarbon stream containing C₈ aromatic hydrocarbons, suchas a reformate stream (product stream of a reformate splitting tower), ahydrocracking product stream, a xylene or EB reaction product stream, anaromatic alkylation product stream, an aromatic disproportionationstream, an aromatic transalkylation stream, a methanol to aromaticproduct stream, and/or a Cyclar™ process stream.

In one embodiment, the feed is the product of the alkylation of benzeneand/or toluene with methanol and/or dimethyl ether in a methylationreactor. One such methylation reactor is described in U.S. Pat. Nos.6,423,879 and 6,504,072, the entire contents of which are incorporatedherein by reference, and employs a catalyst comprising a porouscrystalline material having a Diffusion Parameter for 2,2 dimethylbutaneof about 0.1-15 sec⁻¹ when measured at a temperature of 120° C. and a2,2 dimethylbutane pressure of 60 torr (8 kPa). The porous crystallinematerial may be a medium-pore zeolite, such as ZSM-5, which has beenseverely steamed at a temperature of at least 950° C. in the presence ofat least one oxide modifier, for example, including phosphorus, tocontrol reduction of the micropore volume of the material during thesteaming step. Such a methylation reactor is hereinafter termed a“PX-selective methylation reactor”.

The feed to the PX recovery section may further comprise recyclestream(s) from the isomerization step(s) and/or various separatingsteps. The hydrocarbon feed comprises PX, together with MX, OX, and EB.In addition to xylenes and EB, the hydrocarbon feedstock may alsocontain certain amounts of other aromatic or even non-aromaticcompounds. Examples of such aromatic compounds are C⁷⁻ hydrocarbons,such as benzene and toluene, and C₉₊ aromatics, such as mesitylene,pseudo-cumene and others. These types of feedstream(s) are described in“Handbook of Petroleum Refining Processes”, Eds. Robert A. Meyers,McGraw-Hill Book Company, Second Edition.

The first hydrocarbon feed 102 is initially supplied to a firstsimulated moving bed adsorption unit 120 to recover a first PX-richextract stream 122, an EB-rich extract stream 124, and a PX-depletedraffinate stream 126. Depending on the composition of the hydrocarbonfeed 102, one or more initial separation steps that serve to remove C⁷⁻and C₉₊ hydrocarbons from the feed may occur. Generally the initialseparation steps may include fractional distillation, crystallization,adsorption, a reactive separation, a membrane separation, extraction, orany combination thereof. In embodiments, shown in FIGS. 2 and 3, beforegoing to the first simulated moving bed unit 120, the first hydrocarbonfeed 102 is passed through a xylenes fractionation tower 110 prior topassing to the PX recovery unit 120. The xylenes fractionation tower 110produces an overhead stream 112 comprising C₈ hydrocarbons and a bottomsstream 114 containing C₉₊ hydrocarbons. The overhead stream 112comprising C₈ hydrocarbons is sent to the first PX recovery unit 120,and the bottoms stream 114 containing C₉₊ hydrocarbons may be furtherprocessed, such as in a transalkylation process, to produce xylenes orsent to the motor gasoline pool for use in fuels.

FIG. 4A illustrates a standard simulated moving bed apparatus with 24adsorbent beds. This 24 bed configuration is particularly useful forseparating one C₈ aromatic, such as PX, from a mixture of C₈ aromatics,such as a mixture of PX, MX, OX, and EB. A simulated moving bedadsorption unit uses a solid adsorbent which preferentially retains thePX in order to separate the PX from the rest of the mixture. The solidadsorbent is in the form of a simulated moving bed, where the bed ofsolid adsorbent is held stationary, and the locations at which thevarious streams enter and leave the bed are periodically moved. Theadsorbent bed itself is usually a succession of fixed sub-beds. Theshift in the locations of liquid input and output in the direction ofthe fluid flow through the bed simulates the movement of the solidadsorbent in the opposite direction.

In FIG. 4A, twelve adsorbent beds 401-412 are stacked in a first column491 and another twelve adsorbent beds 413-424 are stacked in a secondcolumn 492. Conduits in fluid communication with a fluid distributiondevice are depicted by double arrows 431-454. The double arrows reflectthe possibility of fluid flow either into or out of columns 491 and 492during the multiple steps of the simulated moving bed process. Forsimplicity, the fluid distribution device is not shown in FIG. 4A. Also,not shown in FIG. 4A are fluid collection areas between beds. However,it will be understood that such collection areas, such as thoserepresented as downcomers as described in U.S. Pat. No. 3,201,491, maybe present in columns 491 and 492 of FIG. 1.

A circulating bulk fluid, which is taken from the bottom of column 492and bed 424, is introduced into the top of column 491 and bed 401through line 462 shown in FIG. 4A. The circulating bulk fluid flows in adownward direction through each of the beds of the first column 491 andis then transported to the top of the second column 492 through line461. The circulating bulk fluid then flows in a downward directionthrough each of the beds of the second column 492.

FIG. 4B shows the flow of fluids through columns 491 and 492 during asingle step of an adsorption cycle in accordance with at least someembodiments disclosed herein. The flow of fluids in FIG. 4B represents astandard simulated moving bed operation, where two extract streams arewithdrawn. In particular, a PX-rich extract stream and an EB-richextract stream is separated from a mixture comprising PX, MX, OX, andEB.

Numbered features in FIG. 4B correspond to numbered features in FIG. 4A.In FIG. 4B, the double arrows in FIG. 4A are replaced with single arrowsto show the actual direction of flow of fluids during a single step inaccordance with at least some embodiments disclosed herein.

The following steps occur at the same time in columns 491 and 492.Overhead stream 112 comprising C₈ hydrocarbons, which comprises amixture of PX, MX, OX, and EB, is introduced into the top of bed 401 incolumn 491 via conduit 431. A first PX-depleted raffinate stream (e.g.,stream 126 is FIGS. 2 and 3), which comprises a desorbent, MX, OX, andEB, is withdrawn from the top of bed 407 through conduit 437. Adesorbent stream is introduced into the top of bed 410 through conduit440. The desorbent may be, for example, PDEB, toluene, or tetralin. Afirst PX-rich extract stream (e.g., stream 122 in FIGS. 2 and 3), whichcomprises desorbent and PX, is withdrawn from the top of bed 416 throughconduit 446, and an EB-rich extract stream (e.g., stream 124 in FIGS. 2and 3), which comprises desorbent, EB, PX, MX, and OX is withdrawn fromthe top of bed 420 through conduit 450.

With respect to the direction of circulating bulk fluid (downwardthrough column 491 and then downward through column 492), both thePX-rich extract stream and EB-rich extract stream are withdrawn atlocations downstream of the desorbent introduction location and upstreamof the feed introduction location. The EB-rich extract stream iswithdrawn downstream of the PX-rich extract stream. The order of the netstreams in this configuration (as shown in FIG. 4B) is feed introductionlocation, raffinate withdrawal location, desorbent introductionlocation, PX-rich extract withdrawal location, and EB-rich extractwithdrawal location. The withdrawal locations for the PX-rich extractstream and EB-rich extract stream may be determined by one skilled inthe art based on the adsorption profile. Generally the withdrawallocation for the PX-rich extract stream is chosen to obtain a purity ofat least 80 wt % PX, preferably at least 90 wt % PX, more preferably atleast 95 wt % PX, and ideally at least 97 wt % PX, based on the totalweight of the PX-rich product stream. As used herein, “purity” of astream means the weight percentage of a certain component in the stream,excluding the desorbent. The EB-rich extract stream should be withdrawnat a location where the recovery of EB is at least 5 wt %, preferably atleast 10 wt %, and more preferably at least 15 wt %, based on the totalweight of EB in the hydrocarbon feed stream. As used herein, “recovery”of a component means the weight percentage of a certain component, basedon the total amount of the component introduced into the system, in thehydrocarbon feed stream.

Referring again to FIGS. 2 and 3, the PX-depleted raffinate stream 126may be separated in fractionation column 130 to produce PX-depletedstream 132 and desorbent stream 134. The desorbent stream 134 isrecycled to the first PX recovery unit 120 and the PX-depleted stream132 may be isomerized to produce more PX.

Because a significant amount of the EB from the first hydrocarbon feed102 has been removed (at least enough to limit buildup in theisomerization loop) as a second extract stream (EB-rich extract stream124) in the first simulated moving bed adsorption unit 120, at least aportion of, and preferably the entirety of, the PX-depleted stream 132is sent to liquid phase isomerization unit 140. Avoiding vapor phaseisomerization saves energy and capital, as liquid phase isomerizationrequires less energy and capital than the vapor phase isomerizationprocess due to the requirement of vaporizing the PX-depleted stream andthe use of hydrogen, which requires an energy- and capital-intensivehydrogen recycle loop. Liquid phase isomerization also reduces thedestruction of xylene molecules into less valuable non-aromatics.

In the liquid phase xylene isomerization unit 140, the PX-depletedstream 132 is contacted with a xylene isomerization catalyst under atleast partially liquid phase conditions effective to isomerize thePX-depleted stream 132 back towards an equilibrium concentration of thexylene isomers. Suitable conditions for the liquid phase isomerizationinclude a temperature of from about 200° C. to about 540° C., preferablyfrom about 230° C. to about 310° C., and more preferably from about 270°C. to about 300° C., a pressure of from about 0 to 6895 kPa(g),preferably from about 1300 kPa(g) to about 3500 kPa(g), a weight hourlyspace velocity (WHSV) of from 0.5 to 100 hr⁻¹, preferably from 1 to 20hr⁻¹, and more preferably from 1 to 10 hr⁻¹. Generally, the conditionsare selected so that at least 50 wt % of the C₈ aromatics would beexpected to be in the liquid phase.

Any catalyst capable of isomerizing xylenes in the liquid phase can beused in the xylene isomerization unit, but in one embodiment thecatalyst comprises an intermediate pore size zeolite having a ConstraintIndex between 1 and 12. Constraint Index and its method of determinationare described in U.S. Pat. No. 4,016,218, which is incorporated hereinby reference. Particular examples of suitable intermediate pore sizezeolites include ZSM-5, ZSM-11, ZSM-12, ZSM-22, ZSM-23, ZSM-35, ZSM-48,and MCM-22, with ZSM-5 and ZSM-11 being particularly preferred,specifically ZSM-5. It is preferred that the acidity of the zeolite,expressed as its alpha value, be greater than 300, such as greater than500, or greater than 1000. The alpha test is described in U.S. Pat. No.3,354,078; in the Journal of Catalysis, Vol. 4, p. 527 (1965); Vol. 6,p. 278 (1966); and Vol. 61, p. 395 (1980), each incorporated herein byreference as to that description. The experimental conditions of thetest used to determine the alpha values cited herein include a constanttemperature of 538° C. and a variable flow rate as described in detailin the Journal of Catalysis, Vol. 61, p. 395. A preferred catalyst isdescribed in U.S. Pat. No. 8,569,559, which is incorporated herein byreference.

The product of the liquid phase xylene isomerization process 140 is afirst isomerized stream 142 having a higher PX concentration than thePX-depleted stream 132. The first isomerized stream 142 is then recycledto the first simulated moving bed adsorption unit 120 to recoveradditional PX and the process is repeated to generate a so-called xyleneisomerization loop. Conducting xylenes isomerization under liquid phaseconditions produces less C₉₊ aromatics than xylenes isomerization undervapor phase conditions. Therefore, the isomerized stream 152 may beprovided to the xylenes fractionation tower 110 at a higher traylocation than an isomerized stream produced by vapor phaseisomerization, yielding greater energy savings. Furthermore, asignificant portion of the isomerized stream 142 may bypass the xylenesfractionation tower 110 and directly enter the first simulated movingbed adsorption unit 120, thereby saving energy by avoiding there-fractionation altogether. Additionally, liquid phase isomerizationreduces the amount of C⁷⁻ hydrocarbons produced in the xyleneisomerization loop, allowing for reduced use of a C⁷⁻ distillation tower(unit 53 in FIG. 1), resulting in further energy savings. Exclusive useof liquid phase isomerization allows for elimination of the C⁷⁻distillation tower as shown in FIG. 2 and FIG. 3.

The first PX-rich extract stream 122 may be separated in a first extractcolumn 150 to remove desorbent and produce a desorbent-free PX-richextract stream 152. As used herein, “desorbent-free” means that thestream contains less than about 1 wt % desorbent, preferably less than0.5 wt % desorbent. The desorbent stream 154 may be recycled to thefirst simulated moving bed adsorption unit 120. The desorbent-freePX-rich stream 152 may be further purified in a first finishing column160 to remove contaminants such as toluene and water 164 and produce afirst pure PX stream 162, which contains at least about 99.7 wt % PX.

The destination for the EB-rich extract stream 124 is dependent upon theamount of PX contained in the stream. Ideally, the withdrawal locationfor the EB-rich extract stream 124 is chosen to minimize the amount ofPX contained in the EB-rich extract stream 124. If the amount of PX inthe EB-rich extract stream 124 is sufficiently low, the EB-rich extractstream 124 may be sent to further processing, such as styreneproduction, or to the motor gasoline pool. While some PX may besacrificed in this embodiment, the energy and cost benefits of nothaving a second separation step to further recover PX may outweigh thedetriment of losing a minimal amount of PX. However, practically it maynot be possible to withdraw the EB-rich extract stream 124 without asignificant amount of PX, and the EB-rich extract stream 124 may be sentto a PX recovery unit, such as a second simulated moving bed adsorptionunit or crystallization unit, to recover a second PX-rich extractstream. Optionally, the desorbent may be removed from the EB-richextract stream 124 prior to the PX recovery unit.

In some embodiments, the first extract column 150 is a dividing wallcolumn that accepts both the PX-rich extract stream 122 and the EB-richextract stream 124. This dividing wall column may remove desorbent fromboth streams while keeping the light components of both feed streamsseparate. In addition, the dividing wall column in these embodiments mayproduce a common heavy product made up mostly of desorbent.

In the embodiment shown in FIG. 2, the EB-rich extract stream 124 issent to a second simulated moving bed adsorption unit 220 to recover thePX in the EB-rich extract stream 124. The EB-rich extract stream 124 maycomprise the entirety of the feed stream to the second simulated movingbed adsorption unit 220 or it may be combined with a second hydrocarbonfeed stream 202 prior to the second simulated moving bed adsorption unit220. The second hydrocarbon feed stream 202 may be the same as the firsthydrocarbon feed stream 102 or be another hydrocarbon stream with thecharacteristics as described above in relation to the first hydrocarbonfeed stream 102. Depending on the composition of the hydrocarbon feed102, one or more initial separation steps that serve to remove C⁷⁻ andC₉₊ hydrocarbons from the feed may occur. Generally the initialseparation steps may include fractional distillation, crystallization,adsorption, a reactive separation, a membrane separation, extractions orany combination thereof. In embodiments, shown in FIG. 2, before goingto the second simulated moving bed 220, the second hydrocarbon feed 202is passed through a xylenes fractionation tower 210 prior to passing tothe PX recovery unit 220. The xylenes fractionation tower 210 producesan overhead stream 212 comprising C₈ hydrocarbons and a bottoms stream214 containing C₉₊ hydrocarbons. The overhead stream 212 comprising C₈hydrocarbons is sent to the second PX recovery unit 220, and the bottomsstream 214 containing C₉₊ hydrocarbons may be further processed, such asin a transalkylation process, to produce xylenes or sent to the motorgasoline pool for use in fuels.

The second simulated moving bed adsorption unit 220 produces a secondPX-rich extract stream 222 and a PX-depleted raffinate stream 224. Thesecond simulated moving bed adsorption unit 220 may contain the sameadsorbent as the first simulated moving bed adsorption unit 120, or itmay contain a different adsorbent than the second simulated moving bedadsorption unit 120 that provides better separation between EB and PX.In particular, in at least some embodiments the adsorbent of the secondsimulated moving bed adsorption unit 220 may have a higher affinity forEB than PX. The second simulated moving bed adsorption unit 220 may be asecond unit such as that shown in FIG. 4A, or a conventional simulatedmoving bed adsorption unit having 24 adsorbent beds divided into twocolumns of 12 adsorbent beds each, such as that shown in FIG. 4A, may beretrofitted such that the first column of 12 adsorbent beds functions asthe first simulated moving bed unit 120 and the second column of 12adsorbent beds functions as the second simulated moving bed unit 220.

The second PX-rich extract stream 222 may be separated in a secondextract column 250 to remove desorbent and produce a desorbent-freePX-rich extract stream 252. The desorbent stream 254 may be recycled tothe second simulated moving bed adsorption unit 220. The desorbent-freePX-rich stream 252 may be further purified in a second finishing column260 to remove contaminants such as toluene and water 264 and produce asecond pure PX stream 262, which contains at least about 99.7 wt % PX.In another embodiment (such as the embodiment of FIG. 7), the secondPX-rich extract stream 222 is provided to the first extract column 150and subsequently to the first finishing column 160, eliminating the needfor a second extract column 250 and second finishing column 260. Inaddition, in the embodiment of FIG. 7 vapor phase isomerization unit240, xylenes fractionation tower 210, fractionation column 270(discussed below) are each also omitted; however, it should beappreciated that these components may be included in other embodimentsthat route PX-rich extract stream 222 to the first extract column asshown in FIG. 7.

Referring again to FIG. 2, the PX-depleted raffinate stream 224 may beseparated in fractionation column 230 to produce PX-depleted stream 232and desorbent stream 234. The desorbent stream 234 is recycled to thesecond PX recovery unit 220 and the PX-depleted stream 232 may beisomerized to produce more PX. Because the PX-depleted raffinate stream224 contains a significant amount of EB, vapor phase isomerization mustbe used in some embodiments to convert the EB therein to limit a buildupof EB in the process. At least a portion of, and preferably the entiretyof, the PX-depleted stream 232 then passes to vapor phase isomerizationunit 240. In other embodiments (e.g., the embodiment of FIG. 7), thePX-depleted raffinate stream 232, which includes a significant amount ofEB is simply emitted from the process and is either sold or utilized inother processes as an EB-containing stream.

In the vapor phase xylene isomerization unit 240, the PX-depleted stream232 is contacted with a xylene isomerization catalyst under at leastpartially vapor phase conditions effective to isomerize the PX-depletedstream 232 back towards an equilibrium concentration of the xyleneisomers. There are generally two types of vapor phase isomerizationcatalysts—one that dealkylates EB to produce benzene and ethylene andisomerizes the xylene isomers, and one that isomerizes the fourdifferent C8 aromatic compounds, including EB, to their equilibriumconcentrations. Either catalyst may be used for the vapor phaseisomerization unit 240.

In one embodiment, the PX-depleted stream 232 is subjected to xylenesisomerization in which the EB in the stream can be dealkylated toproduce benzene. In this embodiment, where the EB is removed bycracking/disproportionation, the PX-depleted C₈ stream is convenientlyfed to a multi-bed reactor comprising at least a first bed containing anEB conversion catalyst and a second bed downstream of the first bed andcontaining a xylene isomerization catalyst. The beds can be in the sameor different reactors. Alternatively, the EB conversion catalyst andxylene isomerization catalyst may be contained in a single bed reactor.

The EB conversion catalyst typically comprises an intermediate pore sizezeolite having a Constraint Index ranging from 1 to 12, a silica toalumina molar ratio of at least about 5, such as at least about 12, forexample at least 20, and an alpha value of at least 5, such as 75 to5000. Constraint Index and its method of determination are disclosed inU.S. Pat. No. 4,016,218, which is herein incorporated by reference,whereas the alpha test is described in U.S. Pat. No. 3,354,078 and inthe Journal of Catalysis, Vol. 4, p. 527 (1965); Vol. 6, p. 278 (1966);and Vol. 61, p. 395 (1980), each incorporated herein by reference as tothat description. The experimental conditions of the test used hereininclude a constant temperature of 538° C. and a variable flow rate asdescribed in detail in the Journal of Catalysis, Vol. 61, p. 395. Higheralpha values correspond with a more active cracking catalyst.

Examples of suitable intermediate pore size zeolites include ZSM-5 (U.S.Pat. Nos. 3,702,886 and Re. 29,948); ZSM-11 (U.S. Pat. No. 3,709,979);ZSM-12 (U.S. Pat. No. 3,832,449); ZSM-22 (U.S. Pat. No. 4,556,477);ZSM-23 (U.S. Pat. No. 4,076,842); ZSM-35 (U.S. Pat. No. 4,016,245);ZSM-48 (U.S. Pat. No. 4,397,827); ZSM-57; ZSM-58; EU-1; and mordenite.The entire contents of the above references are incorporated byreference herein. Preferred zeolites are ZSM-5, ZSM-12 or EU-1.

The zeolite employed in EB conversion catalyst typically has a crystalsize of at least 0.2 microns and exhibits an equilibrium sorptioncapacity for xylene, which can be either para, meta, ortho, or a mixturethereof, of at least 1 gram per 100 grams of zeolite measured at 120° C.and a xylene pressure of 4.5±0.8 mm of mercury and an OX sorption timefor 30 percent of its equilibrium OX sorption capacity of greater than1200 minutes (at the same conditions of temperature and pressure). Thesorption measurements may be carried out gravimetrically in a thermalbalance. The sorption test is described in U.S. Pat. Nos. 4,117,026;4,159,282; 5,173,461; and Re. 31,782, each of which is incorporated byreference herein.

The zeolite used in the EB conversion catalyst may be self-bound (nobinder) or may be composited with an inorganic oxide binder, with thezeolite content ranging from between about 1 to about 99 percent byweight and more usually in the range of about 10 to about 80 percent byweight of the dry composite, e.g., about 65% zeolite with about 35%binder. Where a binder is used, it is preferably non-acidic, such assilica. Procedures for preparing silica bound ZSM-5 are described inU.S. Pat. Nos. 4,582,815; 5,053,374; and 5,182,242, incorporated byreference herein.

In addition, the EB conversion catalyst typically comprises from about0.001 to about 10 percent by weight, e.g., from about 0.05 to about 5percent by weight, e.g., from about 0.1 to about 2 percent by weight ofa hydrogenation/dehydrogenation component. Examples of such componentsinclude the oxide, hydroxide, sulfide, or free metal (i.e., zero valent)forms of Group VIIIA metals (i.e., Pt, Pd, Ir, Rh, Os, Ru, Ni, Co, andFe), Group VIIA metals (i.e., Mn, Tc, and Re), Group VIA metals (i.e.,Cr, Mo, and W), Group VB metals (i.e., Sb and Bi), Group IVB metals(i.e., Sn and Pb), Group IIIB metals (i.e., Ga and In), and Group IBmetals (i.e., Cu, Ag and Au). Noble metals (i.e., Pt, Pd, Ir, Rh, Os andRu) are preferred hydrogenation/dehydrogenation components. Combinationsof catalytic forms of such noble or non-noble metal, such ascombinations of Pt with Sn, may be used. The metal may be in a reducedvalence state, e.g., when this component is in the form of an oxide orhydroxide. The reduced valence state of this metal may be attained, insitu, during the course of a reaction, when a reducing agent, such ashydrogen, is included in the feed to the reaction.

The xylene isomerization catalyst employed in this embodiment typicallycomprises an intermediate pore size zeolite, e.g., one having aConstraint Index between 1 and 12, specifically ZSM-5. The acidity ofthe ZSM-5 of this catalyst, expressed as the alpha value, is generallyless than about 150, such as less than about 100, for example from about5 to about 25. Such reduced alpha values can be obtained by steaming.The zeolite typically has a crystal size less than 0.2 micron and an OXsorption time such that it requires less than 50 minutes to sorb OX inan amount equal to 30% of its equilibrium sorption capacity for OX at120° C. and a xylene pressure of 4.5+0.8 mm of mercury. The xyleneisomerization catalyst may be self-bound form (no binder) or may becomposited with an inorganic oxide binder, such as alumina. In addition,the xylene isomerization catalyst may contain the samehydrogenation/dehydrogenation component as the EB conversion catalyst.

Using the catalyst system described above, EBcracking/disproportionation and xylene isomerization are typicallyeffected at conditions including a temperature of from about 400° F. toabout 1,000° F. (204 to 538° C.), a pressure of from about 0 to about1,000 psig (100 to 7,000 kPa), a weight hourly space velocity (WHSV) ofbetween about 0.1 and about 200 hr−1, and a hydrogen, H₂ to hydrocarbon,HC, molar ratio of between about 0.1 and about 10. Alternatively, theconversion conditions may include a temperature of from about 650° F.and about 900° F. (343 to 482° C.), a pressure from about 50 and about400 psig (446 to 2,859 kPa), a WHSV of between about 3 and about 50 hr−1and a H₂ to HC molar ratio of between about 0.5 and about 5. The WHSV isbased on the weight of catalyst composition, i.e., the total weight ofactive catalyst plus, if used, binder therefor.

In another embodiment, the PX-depleted stream 232 is subjected to EBisomerization to produce a stream containing the C₈ aromatic compoundsin equilibrium concentrations.

Typically, the EB isomerization catalyst comprises an intermediate poresize molecular sieve having a Constraint Index within the approximaterange of 1 to 12, such as ZSM-5 (U.S. Pat. No. 3,702,886 and Re.29,948); ZSM-11 (U.S. Pat. No. 3,709,979); ZSM-12 (U.S. Pat. No.3,832,449); ZSM-22 (U.S. Pat. No. 4,556,477); ZSM-23 (U.S. Pat. No.4,076,842); ZSM-35 (U.S. Pat. No. 4,016,245); ZSM-48 (U.S. Pat. No.4,397,827); ZSM-57; and ZSM-58. Alternatively, the xylene isomerizationcatalyst may comprise a molecular sieve selected from MCM-22 (describedin U.S. Pat. No. 4,954,325); PSH-3 (described in U.S. Pat. No.4,439,409); SSZ-25 (described in U.S. Pat. No. 4,826,667); MCM-36(described in U.S. Pat. No. 5,250,277); MCM-49 (described in U.S. Pat.No. 5,236,575); and MCM-56 (described in U.S. Pat. No. 5,362,697). Themolecular sieve may also comprise a EUO structural type molecular sieve,with EU-1 being preferred, or mordenite. A preferred molecular sieve isone of the EUO structural type having a Si/Al ratio of about 10-25, asdisclosed in U.S. Pat. No. 7,893,309. The entire contents of the abovereferences are incorporated by reference herein.

It may be desirable to combine the molecular sieve of the xyleneisomerization catalyst with another material resistant to thetemperature and other conditions of the process. Such matrix materialsinclude synthetic or naturally occurring substances as well as inorganicmaterials such as clay, silica, and/or metal oxides (such as titaniumoxide or boron oxide). The metal oxides may be naturally occurring or inthe form of gelatinous precipitates or gels including mixtures of silicaand metal oxides. Naturally occurring clays, which can be compositedwith the molecular sieve, include those of the montmorillonite andkaolin families, which families include the subbentonites and thekaolins commonly known as Dixie, McNamee, Ga., and Florida clays orothers in which the main mineral constituent is halloysite, kaolinite,dickite, nacrite, or anauxite. Such clays can be used in the raw stateas originally mined or initially subjected to calcination, acidtreatment, or chemical modification.

In addition to the foregoing materials, the molecular sieve may becomposited with a porous matrix material, such as alumina, zirconia,silica-alumina, silica-magnesia, silica-zirconia, silica-thoria,silica-berylia, silica-titania, aluminum phosphates, titaniumphosphates, zirconia phosphates, as well as ternary compounds such assilica-alumina-thoria, silica-alumina-zirconia, silica-alumina-magnesia,and silica-magnesia-zirconia. A mixture of these components could alsobe used. The matrix may be in the form of a cogel. The relativeproportions of molecular sieve component and inorganic oxide gel matrixon an anhydrous basis may vary widely with the molecular sieve contentranging from between about 1 to about 99 percent by weight and moreusually in the range of about 10 to about 80 percent by weight of thedry composite.

The EB isomerization catalyst also comprises at least one metal fromGroup VIII of the periodic table of the elements and optionally at leastone metal selected from metals from Groups IIIA, IVA, and VIIB. TheGroup VIII metal present in the catalyst used in the isomerizationprocess is selected from iron, cobalt, nickel, ruthenium, rhodium,palladium, osmium, iridium, and platinum, preferably from the noblemetals and highly preferably from palladium and platinum. Morepreferably, the Group VIII metal is platinum. The metal selected fromgroups IIIA, IVA, and VIIB which are optionally present is selected fromgallium, indium, tin, and rhenium, preferably from indium, tin, andrhenium.

The conditions employed in the EB isomerization process generallyinclude a temperature of from 300 to about 500° C., preferably fromabout 320 to about 450° C. and more preferably from about 340 to about430° C.; a partial pressure of hydrogen from about 0.3 to about 1.5 MPa,preferably from about 0.4 to about 1.2 MPa, and more preferably fromabout 0.7 to about 1.2 MPa; a total pressure of from about 0.45 to about1.9 MPa, preferably from about 0.6 to about 1.5 MPa; and a weight hourlyspace velocity (WHSV) of between about 0.25 and about 30 hr⁻¹,preferably between about 1 and about 10 hr⁻¹ and more preferably betweenabout 2 and about 6 hr⁻¹.

The product of the vapor phase xylene isomerization process 240 is asecond isomerized stream 242 having a higher PX concentration than thesecond PX-depleted raffinate stream 224 or the second PX-depleted stream232. The second isomerized stream 242 is then recycled to one or both ofthe first simulated moving bed adsorption unit 120 the second simulatedmoving bed adsorption unit 220. Prior to the first or second simulatedmoving bed adsorption units 120, 220, the second isomerized stream 242may be sent through fractionation column 270 to remove the C⁷⁻hydrocarbons 274 and produce a C₈₊ hydrocarbon stream 272, which may beprocessed through the xylenes fractionation column 110 or 210.

In an embodiment in which the EB-rich extract stream 124 comprises theentirety of the feed stream to the second simulated moving bedadsorption unit 220 (not shown), the PX-depleted raffinate stream 224may be essentially free of other xylene isomers. Thus, the fractionationcolumn 230 would produce an overhead stream that is essentially pure EB,which would then be available for the motor gasoline pool or otherrefinery processes. In this embodiment, there is no need for anyisomerization in of the raffinate stream of the second simulated movingbed adsorption unit, eliminating the need for vapor phase isomerization.

In the embodiment shown in FIG. 3, the EB-rich extract stream 124 issent to a crystallization unit 300 to recover the PX in the EB-richextract stream 124. The crystallization unit 300 produces a secondPX-rich extract stream 302 and an EB-rich liquid product 304. The secondPX-rich extract stream 302 may be purified using the conventionalcrystallization equipment or may be sent to the first extract column 150and subsequently to the first finishing column 160, eliminating the needfor a second extract column and second finishing column. The EB-richliquid product 304 may be sent to a refinery stream or may be furtherpurified for downstream processes.

EXAMPLES

In the Examples that follow, a computer model is used to simulate theseparation of PX from other C₈ aromatics in the first simulated movingbed unit 120 as shown in FIGS. 2 and 3. The unit comprises two columns,as shown in FIG. 4A, in fluid communication with a rotary valve device.Each column comprises twelve adsorbent bed chambers, stacked one on topof the other, containing an adsorbent. For the purposes of explanation,these beds are identified as beds 101 to 124. The number of bedsdescribed in each zone is for illustrative purposes and the number ofbeds may be varied without changing the concepts described herein.

In the first column, the beds are stacked such that fluid introducedinto the top of the first column flows downward through the first bed(i.e., bed 101) and then through the beds below to the last bed (i.e.,bed 112) in the first column. Fluid from the bottom of the first columnthen flows to the top of the second column where it flows downwardthrough the beds (i.e., beds 113-124). Fluid from the bottom of thesecond column then flows to the top of the first column to complete aloop of circulating bulk fluid throughout the columns.

The initial introduction of feed may take place in any of the beds ofthe apparatus. For example, feed may be introduced to the first bed inthe first column. The feed is primarily composed of a mixture of C₈aromatics. The feed may also include small amounts of impuritiesincluding toluene and paraffins. The feed may be a mixture of productstreams from a reforming process, a transalkylation process, and anisomerization process.

When a steady state operation of the simulated moving bed (SMB) unit isachieved, the beds of the apparatus may be described in terms of foursub-zones, i.e., a desorption sub-zone, a purification or rectificationsub-zone, an adsorption sub-zone, and a buffer sub-zone. In a standardSMB unit, (1) the desorption sub-zone may include the bed to which adesorbent stream is introduced and six beds downstream from this bedterminating in the bed from which the extract stream is withdrawn, (2)the purification sub-zone may include nine beds immediately downstreamof desorption sub-zone, terminating with the bed immediately upstreamfrom the bed to which feed is introduced, (3) the adsorption sub-zonemay include the bed to which feed is introduced and six beds immediatelydownstream of the purification sub-zone terminating in the bed fromwhich a raffinate stream is withdrawn, and (4) the buffer sub-zone mayinclude three beds immediately downstream from the purification sub-zoneand terminating in the bed immediately upstream from the desorptionsub-zone. The number of beds in each zone may vary from the numbersdescribed above.

The raffinate and extract streams may pass through conduits and througha rotary valve device. These streams may then be distilled to separatedesorbent from C₈ aromatics. A PX product may be recovered from thedistillation of the extract stream. MX, OX, and EB obtained bydistillation of the raffinate stream may be passed to an isomerizationunit to convert a portion of these C₈ aromatics to PX, and theisomerized C₈ aromatics may then be used as a portion of the feed to theadsorption process. Desorbent recovered by distillation of the extractand raffinate streams may be recycled to the adsorption process.

Parameters for Examples 1 and 2

In Examples 1 and 2 which follow, an SMB model that consists of 24 beds(length 1.135 m, cross-sectional area 13.3 m²) was employed. A mixtureof C₈ aromatics (PX, OX, MX, and EB) and desorbent (PDEB) was assumed tobe fed to the unit. Two extract streams, a PX-rich extract stream,herein referred to as Extract1, and an EB-rich extract stream, hereinreferred to as Extract2, were withdrawn at locations between theintroduction locations of the desorbent and the feed, i.e., in thepurification zone.

The zone configuration is consistent in this study. For an SMB with 24beds, the zone configuration is fixed to 6:4:5:6:3 (i.e., six bedsbetween desorbent and Extract1, four beds between Extract1 and Extract2,five beds between Extract2 and feed, six beds between feed andraffinate, and three bed between raffinate and desorbent).

Consistent with M. Minceva, and A. E. Rodrigues, ‘Modeling andSimulation of a Simulated Moving Bed for the Separation of P-Xylene’,Industrial & Engineering Chemistry Research, 41 (2002), 3454-61, thefollowing assumptions are made: (1) isothermal, isobaric operation; (2)constant velocity within each zone; (3) solid phase concentration ishomogeneous throughout adsorbent particles; and (4) the mass transferbetween the liquid and adsorbent phases is described by the lineardriving force (LDF) model.

Based on these assumptions, mass balance equations can be written as:

∂ C ik ⁡ ( z , t ) ∂ t = LK ⁢ ( t ) ⁢ ∂ 2 ⁢ C ik ⁡ ( z , t ) ∂ z 2 - v k * ⁡ (t ) ⁢ ∂ C ik ⁡ ( z , t ) ∂ z - ( 1 - ɛ ) ɛ ⁢ ∂ q ik ⁡ ( z , t ) ∂ twhere i is the index for components (i=PX, MX, OX, EB, PDEB); k is theindex for columns (k=1 . . . N_(bed), where N_(bed) is the total numberof beds); C is the bulk liquid concentration

$( {{unit}\;\frac{kg}{m^{3}}} );$q is the adsorbate concentration

$( {{unit}\;\frac{kg}{m^{3}}} );$ε is the overall porosity;

is the axial dispersion coefficient; and v_(k)* is the interstitialvelocity in columns.

This mass balance equation describes the change of bulk liquidconcentration at a specific position inside of a column (first term)with respect to dispersion (second term), convection (third term), andadsorption/desorption process (fourth term).

The LDF model is written as:

$\frac{\partial{q_{ik}( {z,k} )}}{\partial t} = {k( {{q_{ik}^{*}( {z,t} )} - {q_{ik}( {z,t} )}} )}$where q* is the adsorbate concentration in equilibrium with the liquidphase

$( {{unit}\;\frac{kg}{m^{3}}} ).$The LDF model describes the mass flux into the solid phase. Theadsorbate concentration in equilibrium with the liquid phase can beobtained from an adsorption isotherm.

At the node between columns, the mass balance is calculated bysubtracting outlet flow rates and adding inlet flow rates:F _(k+1) =F _(k) +F _(Feed,k) +F _(desorbent,k) −F _(raffinate,k) −F_(extract1,k) −F _(extract2,k)

For columns that are not connected to inlet or outlet streams,F_(Feed,k) or F_(desorbent,k) or F_(raffinate,k) or F_(extract1,k) orF_(extract2,k) is zero.

The CSS constraints are given as:C _(k+1)(z,t _(end))=C _(k)(z,t ₀)where t_(end) is the time at the end of a step, and to is the beginningof a step. Here, stepwise symmetry is assumed, where every step isidentical.

Model parameters were taken from the literature, in particular, from M.Minceva, and A. E. Rodrigues, ‘Modeling and Simulation of a SimulatedMoving Bed for the Separation of P-Xylene’, Industrial & EngineeringChemistry Research, 41 (2002), 3454-61. Model parameters are summarizedin Table 1.

TABLE 1 SMB unit Geometry Model Parameter L_(c) = 113.5 cm P_(e) =ν_(k)L_(k)/D_(Lk) = 2000 d_(c) = 411.7 cm k = 2 min⁻¹ V_(c) = 15.1 × 10⁶cm³ d_(p) = 0.092 cm No. of Columns = 24 ε = 0.39 Configuration =6-9-6-3 ρ = 1.39 g/cm³ q_(mPX(MX;OX;EB)) = 130.3 mg/g K_(PX) = 1.0658cm³/mg K_(MX) = 0.2299 cm³/mg K_(OX) = 0.1884 cm³/mg K_(EB) = 0.3067cm³/mg q_(mPDEB) = 107.7 mg/g K_(PDEB) = 1.2935 cm³/mg

The mass transfer coefficient was changed from 2 min⁻¹ to 0.75 min⁻¹.

The optimization problem was formulated as follows:

-   -   Objective function: maximize F_(Feed)    -   Decision variables: F₁, F₂, F₃, F₄, F₅, t_(st)        -   where F_(j)'s are zone flow rates, and t_(st) is the step            time    -   Main Constraint: Extract1 purity (PX)≥99.7%        -   Extract(1+2) recovery (PX)≥97.0%

The model was discretized into a set of algebraic differential equationsby applying the center finite difference method (CFDM) to the spatialdomain and orthogonal collocation finite element method (OCFEM) to thetemporal domain respectively. The discretized problem was solved by aninterior-point optimization algorithm, IPOPT.

Example 1

In this Example, additional constraints were imposed to ensure 15 wt %of the total amount of EB fed into the SMB system was recovered in theEB-rich extract stream (Extract2), and less than 1.0 wt % of the totalOX and MX in the feed were recovered in Extract2. At least 15 wt % EBrecovery in Extract2 minimizes EB build up in the xylene loop usingliquid phase isomerization, allowing for significant energy savings byeliminating vapor phase isomerization. The corresponding EB compositionin the feed to the SMB system increases to 14 wt % due to EB recyclefrom the isomerization units.

FIG. 5 shows the concentration profile within the SMB system at thebeginning of a step. The system is divided into five zones based on thetwo inlet streams and three outlet streams. The locations of the feedand desorbent inlets and raffinate and extract outlets are shown witharrows. The PX-rich extract stream (Extract1) and EB-rich extract stream(Extract2) are withdrawn from the upstream of the feed location whilethe raffinate stream is withdrawn from the downstream side of the feedlocation.

Table 2 shows the optimized mass flow rates of each component in theproduct outlets. The amount of PDEB has been unevenly split among theoutlet streams. The EB-rich extract (Extract2) recovers only 1.3 wt % ofthe total PDEB in the SMB system. Thus, the PDEB may not need to beremoved from the EB-rich extract (Extract2) prior to the PX recoveryunit, eliminating the requirement of an additional fractionating tower.

TABLE 2 OX PX MX EB PDEB Feed (kg/min) 241.5 448.8 945.2 266.2 0Desorbent (kg/min) 0 0 0 0 1543 Raffinate (kg/min) 241 13.5 934.5 225.31060.6 Extract 1 (kg/min) 0.01 409 0.04 0.75 464 Extract 2 (kg/min) 0.2227.8 9.43 39.8 19.3

Table 3 shows the purity and recovery obtained in all of the productoutlets. The PX-rich extract stream (Extract1) recovers 91.1 wt % of PXwith 99.8 wt % purity. The EB-rich extract stream (Extract2) recovers anadditional 6.2 wt % of PX to ensure an overall recovery of PX of morethan 97 wt %. The purity values of PX and EB in the EB-rich extractstream (Extract2) are 28.8 wt % and 41.3 wt %. This EB-rich extractstream (Extract2) can be either fed to a downstream SMB or acrystallizer to obtain pure PX and EB fractions. The desorbent to feedratio required in this example is 0.81, which is less than required in aconventional Parex system, meaning the benefits of this multiple-extractscheme may come without increasing the energy consumption.

TABLE 3 OX PX MX EB Raffinate purity (wt %) 17 0.95 66.1 15.9 Raffinaterecovery (wt %) 99.8 3 98.9 84.6 Extract 1 purity (wt %) 0 99.8 0.010.18 Extract 1 recovery (wt %) 0 91.1 0 0.28 Extract 2 purity (w/PDEB)(wt %) 0.22 28.8 9.8 41.3 Extract 2 recovery (w/PDEB) (wt %) 0.1 6.2 115

Example 2

In this Example, additional constraints were imposed to ensure 25 wt %of the total amount of EB fed into the SMB system was recovered in theEB-rich extract stream (Extract2), and less than 2.5 wt % of the totalOX and MX in the feed were recovered in Extract2. At least 15 wt % EBrecovery in Extract2 minimizes EB build up in the xylene loop usingliquid phase isomerization, allowing for significant energy savings byeliminating vapor phase isomerization. The corresponding EB compositionin the feed to the SMB system increases to 14 wt % due to EB recyclefrom the isomerization units.

FIG. 6 shows the concentration profile within the SMB system at thebeginning of a step. The system is divided into five zones based on thetwo inlet streams and three outlet streams. The locations of the feedand desorbent inlets and raffinate and extract outlets are shown witharrows. The PX-rich extract stream (Extract1) and EB-rich extract stream(Extract2) are withdrawn from the upstream of the feed location whilethe raffinate stream is withdrawn from the downstream side of the feedlocation.

Table 4 shows the optimized mass flow rates of each component in theproduct outlets. The amount of PDEB has been unevenly split among theoutlet streams. The EB-rich extract (Extract2) recovers only 3.5 wt % ofthe total PDEB in the SMB system. Thus, the PDEB may not need to beremoved from the EB-rich extract (Extract2) prior to the PX recoveryunit, eliminating the requirement of an additional fractionating tower.

TABLE 4 OX PX MX EB PDEB Feed (kg/min) 164.9 306.4 645.3 181.8 0Desorbent (kg/min) 0 0 0 0 1710.8 Raffinate (kg/min) 164.8 9.2 630.7135.9 978.5 Extract 1 (kg/min) 0.03 264.2 0.1 0.81 672.6 Extract 2(kg/min) 0.49 33.6 16.15 45.5 59

Table 5 shows the purity and recovery obtained in all of the productoutlets. The PX-rich extract stream (Extract1) recovers 86.2 wt % of PXwith 99.8 wt % purity. The EB-rich extract stream (Extract2) recovers anadditional 11 wt % of PX to ensure an overall recovery of PX of morethan 97 wt %, and recovers 25% of the total EB fed into the SMB systemto further minimize EB build up in the xylene loop. The purity values ofPX and EB in the EB-rich extract stream (Extract2) are 21.7 wt % and29.4 wt %. This EB-rich extract stream (Extract2) can be either fed to adownstream SMB or a crystallizer to obtain pure PX and EB fractions.

TABLE 5 OX PX MX EB Raffinate purity (wt %) 17.5 0.98 67 14.5 Raffinaterecovery (wt %) 99.9 3 97.7 74.8 Extract 1 purity (wt %) 0.01 99.7 0.040.3 Extract 1 recovery (wt %) 0.02 86.2 0.02 0.44 Extract 2 purity(w/PDEB) (wt %) 0.32 21.7 10.44 29.4 Extract 2 recovery (w/PDEB) (wt %)0.3 11 2.5 25

Examples 1 and 2 illustrate the flexibility of the embodiments disclosedherein. Both examples describe a means of eliminating vapor phaseisomerization, while creating a concentrated EB product stream. The modeof operation can be varied depending on the equipment limitations in thexylene loop. In Example 1, the recovery of pure PX product in Extract1is higher, but at the cost of reduced EB recovery in the Extract2stream, which will lead to larger recycle streams and a larger desorbentto feed ratio. Example 2 has a higher EB recovery in Extract2, whichwould reduce the size of recycle streams in the xylene loop, but at thecost of reduced recovery of PX in the Extract1 stream.

While various embodiments have been disclosed herein, modificationsthereof can be made without departing from the scope or teachingsherein. In particular, many variations and modifications of the systems,apparatus, and processes described herein are possible and are withinthe scope of the disclosed subject matter. Accordingly, embodimentsdisclosed herein are exemplary only and are not limiting. As a result,the scope of protection is not limited to the embodiments describedherein, but is only limited by the claims that follow, the scope ofwhich shall include all equivalents of the subject matter of the claims.Unless expressly stated otherwise, the steps in a method claim may beperformed in any order. The use of identifiers such as (a), (b), (c)before steps in a method claim is not intended to and does not specify aparticular order to the steps. Rather the use of such identifiers areused to simplify subsequent reference to such steps. Finally, the use ofthe term “including” in both the description and the claims is used inan open ended fashion, and should be interpreted as meaning “including,but not limited to”.

Trade names used herein are indicated by a ™ symbol or ® symbol,indicating that the names may be protected by certain trademark rights,e.g., they may be registered trademarks in various jurisdictions. Allpatents and patent applications, test procedures (such as ASTM methods,UL methods, and the like), and other documents cited herein are fullyincorporated by reference to the extent such disclosure is notinconsistent with this invention and for all jurisdictions in which suchincorporation is permitted. When numerical lower limits and numericalupper limits are listed herein, ranges from any lower limit to any upperlimit are contemplated.

The invention claimed is:
 1. A process for recovering paraxylene, theprocess comprising: (a) introducing a first hydrocarbon feed stream intoa first simulated moving bed adsorption unit, wherein the firsthydrocarbon feed stream comprises a mixture of paraxylene (PX),metaxylene (MX), orthoxylene (OX), and ethylbenzene (EB); (b)introducing a desorbent stream into the first simulated moving bedadsorption unit, wherein the desorbent stream comprises a desorbent; (c)withdrawing a first PX-rich extract stream, which comprises desorbentand PX, from the first simulated moving bed adsorption unit, wherein thefirst PX-rich extract stream is withdrawn at a location downstream of adesorbent introduction location and upstream of a feed introductionlocation; (d) withdrawing an EB-rich extract stream, which comprisesdesorbent, EB, and PX, from the first simulated moving bed adsorptionunit, wherein the EB-rich extract stream is withdrawn at a locationdownstream of the first PX-rich extract stream withdrawal location andupstream of the feed introduction location; (e) withdrawing a firstPX-depleted raffinate stream, which comprises desorbent, MX, OX, and EB,from the first simulated moving bed adsorption unit; (f) isomerizing atleast a portion of the first PX-depleted raffinate stream at leastpartially in a liquid phase to produce a first isomerized stream havinga higher PX concentration than the first PX-depleted raffinate stream;(g) recycling at least a portion of the first isomerized stream to thefirst simulated moving bed adsorption unit; and (h) providing theEB-rich extract stream to a PX recovery unit to produce a second PX-richextract stream, wherein the PX recovery unit comprises a secondsimulated moving bed adsorption unit or a crystallizer.
 2. The processof claim 1, further comprising: (i) providing the first PX-rich extractstream to a first extract column to remove desorbent and produce adesorbent-free PX-rich extract stream; and (j) providing thedesorbent-free PX-rich extract stream to a first finishing column toproduce a first pure PX stream.
 3. The process of claim 2, wherein thesecond PX-rich extract stream is provided to the first extract column ofstep (i).
 4. The process of claim 3, wherein the second simulated movingbed adsorption unit produces a second PX-depleted raffinate stream, andfurther comprising recovering a pure EB stream from the secondPX-depleted raffinate stream in a distillation column.
 5. The process ofclaim 4, further comprising removing desorbent from the EB-rich extractstream before step (h).
 6. The process of claim 4, wherein the secondsimulated moving bed adsorption unit comprises a different adsorbentthan the first simulated moving bed adsorption unit, and wherein theadsorbent of the second moving bed adsorption unit has a higher affinityfor EB and PX than the adsorbent of the first moving bed adsorptionunit.
 7. The process of claim 4, wherein a conventional simulated movingbed adsorption unit comprising 24 adsorbent beds, divided into twocolumns of 12 adsorbent beds each, is retrofitted such that the firstcolumn functions as the first simulated moving bed adsorption unit andthe second column functions as the second simulated moving bedadsorption unit.
 8. The process of claim 4, wherein the EB-rich extractstream is combined with a second hydrocarbon feed stream before step(h).
 9. The process of claim 4, wherein the withdrawal location of theEB-rich extract stream from the first simulated moving bed adsorptionunit is selected to minimize a desorbent content in the EB-rich extractstream, and the EB-rich extract stream comprises the entirety of a feedstream to the second simulated moving bed adsorption unit.
 10. Theprocess of claim 4, wherein the withdrawal location of the EB-richextract stream from the first simulated moving bed adsorption unit is atleast 3 beds away from the first PX-rich extract stream withdrawallocation.
 11. The process of claim 3, wherein the second simulatedmoving bed adsorption unit produces a second PX-depleted raffinatestream, and further comprising: (k) isomerizing at least a portion ofthe second PX-depleted raffinate stream in a vapor phase to produce asecond isomerized stream having a higher PX concentration than thesecond PX-depleted raffinate stream; and (l) recycling at least aportion of the second isomerized stream to one or both of the firstsimulated moving bed adsorption unit and the second simulated moving bedadsorption unit.
 12. The process of claim 3, wherein the secondsimulated moving bed adsorption unit produces an EB-rich raffinatestream and an EB-lean raffinate stream, and further comprising: (m)recycling at least a portion of the EB-lean raffinate stream to step(f); and (o) isomerizing at least a portion of the EB-rich raffinatestream in a vapor phase to produce a second isomerized stream having ahigher PX concentration than the EB-rich raffinate stream.
 13. Theprocess of claim 2, wherein the PX recovery unit further comprises acrystallizer.
 14. The process of claim 13, wherein the second PX-richextract stream is provided to the first extract column of step (i). 15.The process of claim 14, wherein the crystallizer produces an EB-richliquid product, which is sent to a refinery stream.
 16. A process forrecovering paraxylene, the process comprising: (a) providing a firsthydrocarbon feed stream and a desorbent stream to a first simulatedmoving bed adsorption unit, wherein the first hydrocarbon feed streamcomprises a mixture of paraxylene (PX), metaxylene (MX), orthoxylene(OX), and ethylbenzene (EB), and wherein the desorbent stream comprisesdesorbent; (b) withdrawing from the first simulated moving bedadsorption unit: (i) a first PX-rich extract stream, which comprisesdesorbent and PX; (ii) an EB-rich extract stream, which comprisesdesorbent, EB, and PX; and (iii) a first PX-depleted raffinate stream,which comprises desorbent, MX, OX, and EB, wherein the first PX-richextract stream is withdrawn at a location downstream of a desorbentintroduction location and upstream of a feed introduction location andthe EB-rich extract stream is withdrawn at a location downstream of thefirst PX-rich extract stream withdrawal location and upstream of thefeed introduction location; (c) isomerizing at least a portion of thefirst PX-depleted raffinate stream at least partially in a liquid phaseto produce a first isomerized stream having a higher PX concentrationthan the first PX-depleted raffinate stream; (d) recycling at least aportion of the first isomerized stream to the first simulated moving bedadsorption unit; (e) providing the EB-rich extract stream to a secondsimulated moving bed adsorption unit to produce a second PX-rich extractstream and a second PX-depleted raffinate stream; and (f) recovering apure EB stream from the second PX-depleted raffinate stream in adistillation column.
 17. The process of claim 16, further comprisingremoving desorbent from the EB-rich extract stream before step (e). 18.The process of claim 16, wherein the EB-rich extract stream is combinedwith a second hydrocarbon feed stream before step (e).
 19. The processof claim 16, further comprising removing C₉₊ hydrocarbons from the firsthydrocarbon feed stream in a xylenes fractionation column prior to step(a).
 20. A process for recovering paraxylene, the process comprising:(a) providing a first hydrocarbon feed stream and a desorbent stream toa first simulated moving bed adsorption unit, wherein the firsthydrocarbon feed stream comprises a mixture of paraxylene (PX),metaxylene (MX), orthoxylene (OX), and ethylbenzene (EB), and whereinthe desorbent stream comprises desorbent; (b) withdrawing from the firstsimulated moving bed adsorption unit: (i) a first PX-rich extractstream, which comprises desorbent and PX; (ii) an EB-rich extractstream, which comprises desorbent, EB, and PX; and (iii) a firstPX-depleted raffinate stream, which comprises desorbent, MX, OX, and EB,wherein the first PX-rich extract stream is withdrawn at a locationdownstream of a desorbent introduction location and upstream of a feedintroduction location and the EB-rich extract stream is withdrawn at alocation downstream of the first PX-rich extract stream withdrawallocation and upstream of the feed introduction location; (c) isomerizingat least a portion of the first PX-depleted raffinate stream at leastpartially in a liquid phase to produce a first isomerized stream havinga higher PX concentration than the first PX-depleted raffinate stream;(d) recycling at least a portion of the first isomerized stream to thefirst simulated moving bed adsorption unit; (e) providing the EB-richextract stream to a second simulated moving bed adsorption unit toproduce a second PX-rich extract stream and a second PX-depletedraffinate stream; (f) isomerizing at least a portion of the secondPX-depleted raffinate stream in a vapor phase to produce a secondisomerized stream having a higher PX concentration than the secondPX-depleted raffinate stream; and (g) recycling at least a portion ofthe second isomerized stream to the first simulated moving bedadsorption unit or the second simulated moving bed adsorption unit. 21.The process of claim 20, further comprising removing C⁷⁻ hydrocarbonsfrom the second isomerized stream in a fractionation column prior tostep (g).
 22. The process of claim 20, further comprising removing C₉₊hydrocarbons from the second isomerized stream in a xylenesfractionation column prior to step (g).